JP4808209B2 - Process for producing lower olefins and aromatic hydrocarbons - Google Patents

Abstract

A process for producing light olefins and aromatics, which comprises reacting a feedstock by contacting with a catalytic cracking catalyst in at least two reaction zones, wherein the reaction temperature of at least one reaction zone among the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone and its weight hourly space velocity is lower than that of the first reaction zone, separating the spent catalyst from the reaction product vapor, regenerating the separated spent catalyst and returning the regenerated catalyst to the reactor, and separating the reaction product vapor to obtain the desired products, light olefins and aromatics. This process produces maximum light olefins such as propylene, ethylene, etc from heavy feedstocks, wherein the yield of propylene exceeds 20% by weight, and produces aromatics such as toluene, xylene, etc at the same time.

Description

The present invention relates to a process for the catalytic conversion of hydrocarbon oils in the absence of hydrogen, more particularly, it comprises heavy feedstocks as light olefins rich in propylene and ethylene and toluene and xylenes. The present invention relates to a process for conversion to aromatic hydrocarbons rich in water.

Background of the Invention Light olefins such as ethylene, propylene and the like are important chemical raw materials, where propylene is a synthetic monomer for products such as polypropylene, acrylonitrile and the like. Along with the rapid increase in demand for derivatives such as polypropylene, the demand for propylene is also increasing year by year, for example, the global market demand for propylene increases by an average of 6.3% annually. The rate has already increased from 15,200,000 tons 20 years ago to 51,200,000 tons in 2000. The demand for propylene is expected to reach 86,000,000 tons by 2010, with an average annual growth rate of 5.6%.

  Propylene production methods are mainly steam cracking and catalytic cracking (FCC), where steam cracking uses light oil such as naphtha as raw material, with a propylene yield of only about 15% by weight. Ethylene and propylene are produced by pyrolysis, while FCC uses heavy oils such as vacuum gas oil (VGO) as feedstock. Currently, 66% of the world's propylene comes from steam cracking byproducts to produce ethylene, 32% from refinery FCC byproducts to produce naphtha and diesel, A small amount (about 2%) is then obtained from propane dehydrogenation and ethylene and butene metathesis.

  When the petrochemical industry produces ethylene and propylene via conventional steam cracking pathways, it faces several limiting factors such as a lack of light feedstock, a lack of processing capacity and high production costs.

  FCC is gaining more and more attention due to its advantages of wide range adaptability, flexible operation and so on. In the United States, nearly 50% of the market demand for propylene comes from FCC units. The development of improved catalytic cracking technology to increase the production of propylene is very rapid.

U.S. Pat.No. 4,980,053 discloses a conversion process for producing light olefins from hydrocarbons, wherein the feed is a petroleum fraction, residual oil, or crude oil having various boiling ranges, and The conversion reaction is performed by using a solid acid catalyst, a temperature of 500 ° C. to 650 ° C., a pressure of 1.5′10 ^{5 to} 3′10 ⁵ Pa, a WHSV of 0.2 h ^{−1 to} 2.0 h ⁻¹ , And in a fluidized bed or moving bed reactor under conditions of 2 to 12 ratio of catalyst to oil. After the catalyst is regenerated by burning the coke, it returns to the reactor for recycle use. By this process, the total yield of propylene and ethylene can reach about 40%, where the yield of propylene is up to 26.34%.

  WO 00 / 31215A1 discloses a catalytic cracking process for producing olefins, which has ZSM-5 and / or ZSM-11 zeolite as the active ingredient and a large amount of inert material as support. A catalyst is used and VGO is used as a feedstock. The yield of propylene does not exceed 20% by weight.

  U.S. Pat.No. 6,123,830 discloses a combination process consisting of two-stage catalytic cracking and two-stage hydroprocessing, the purpose of which is to produce as many olefins as possible and the quality of the oil distillate and the naphtha octane number. Is to improve. The feedstock is converted into a first hydroprocessing product in a first hydroprocessing unit, and the first hydroprocessing product enters a first catalytic cracking unit, where naphtha, diesel and heavy oil are the main active It is obtained by using a catalyst having intermediate pore size zeolite as a component. Heavy oil enters a second hydroprocessing unit for hydroprocessing to obtain a second hydroprocessing product, and the second hydroprocessing product enters a second catalytic cracking unit for cracking, wherein The active component of the catalyst is mainly an intermediate pore size zeolite. The yield of propylene in this process is quite low.

  Aromatic hydrocarbons are also important chemical raw materials, especially light aromatic hydrocarbons such as BTX (benzene, toluene, and xylene), which produce synthetic materials such as chemical fibers, plastics, etc. Used to do. At present, the main method for producing aromatic hydrocarbons is catalytic reforming, where the active component of the reforming catalyst is a noble metal, so the raw material should be fed to a strict pretreatment. In addition, the process flow of reforming catalyst transfer and regeneration is complex.

  The above prior art produces propylene only as a by-product in low yields not exceeding 30% at the same time that naphtha and diesel are produced. Only some of them can produce aromatic hydrocarbons, but light olefins and aromatic hydrocarbons cannot be produced simultaneously. To meet the increasing demand for chemical raw materials, propylene, ethylene, aromatic hydrocarbons, etc., chemical industry type for simultaneously producing large quantities of propylene, ethylene and aromatic hydrocarbons from heavy feedstocks There is a need to develop an oil refining process.

BRIEF DESCRIPTION OF THE INVENTION The object of the present invention is to provide several processes for the simultaneous production of propylene, ethylene and aromatic hydrocarbons from heavy feeds based on the prior art, and the propylene in these processes. The yield is higher than 20%.

Technical scheme 1
The process for producing light olefins and aromatic hydrocarbons provided by the present invention includes:
The feedstock contacts the catalytic cracking catalyst and reacts under conditions of reaction temperature of 400 ° C. to 800 ° C. and WHSV of 0.1 h ^{−1 to} 750 h ⁻¹ . The reaction is carried out in at least two reaction zones, and the reaction temperature of at least one of the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone; and Its WHSV is lower than that of the first reaction zone. The spent catalyst is separated from the reaction product vapor, and the catalyst returns to the reactor after being regenerated. The reaction product vapor is separated to yield the desired product, light olefins and aromatic hydrocarbons.

Technical scheme 2
Another method for producing light olefins and aromatic hydrocarbons provided by the present invention includes the following:
(1) The raw material and optional cycle material enter the hydroprocessing unit, come into contact with the hydroprocessing catalyst and hydrogen, and the hydrogen partial pressure of 3.0 MPa to 20.0 MPa, the reaction temperature of 300 ° C. to 450 ° C. Reaction under conditions of 300 to 2000 hydrogen / oil volume ratio and 0.1 h ^{−1 to} 3.0 h ⁻¹ LHSV, and then separating the reaction effluent to produce the hydrotreated product. And hydrogen is recycled for reuse;
(2) The hydrotreated product is contacted with a catalytic cracking catalyst and reacted under conditions of a reaction temperature of 400 ° C. to 800 ° C. and a WHSV of 0.1 h ^{−1 to} 750 h ⁻¹ . The reaction is carried out in at least two reaction zones, and the reaction temperature of at least one of the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone; and Its WHSV is lower than that of the first reaction zone. The spent catalyst is separated from the reaction product vapor, and the catalyst returns to the reactor after being regenerated. The reaction product vapor is separated to yield the desired product, light olefins and aromatic hydrocarbons.

Technical scheme 3
The third method for producing light olefins and aromatic hydrocarbons provided by the present invention includes the following:
(1) The raw material is in contact with hydrogen and a hydrotreating catalyst, and a hydrogen partial pressure of 3.0 MPa to 20.0 MPa, a reaction temperature of 300 ° C. to 450 ° C., a hydrogen / oil volume ratio of 300 to 2000, and react under the conditions of LHSV of 1h ^-1 ~3.0h ^-1, and then separating the reaction effluent _{(reaction effluent), H 2,} CH 4, hydrotreated _C ² 0 _-C ⁴ 0, Hydrotreating naphtha and hydrotreating product are obtained and hydrogen is recycled for reuse;
(2) The hydrotreated product of step (1) is contacted with a catalytic cracking catalyst and reacted under conditions of a reaction temperature of 400 ° C. to 800 ° C. and a WHSV of 0.1 h ^{−1 to} 750 h ⁻¹ . The reaction is carried out in at least two reaction zones, and the reaction temperature of at least one of the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone; and Its WHSV is lower than that of the first reaction zone. The spent catalyst is separated from the reaction product vapor and stripped and regenerated before returning to all or part of the reaction zone of step (2). The reaction product vapor is separated, H ₂ , CH ₄ , catalytic cracking C ₂ ⁼ -C ₃ ⁼ , catalytic cracking C ₂ ⁰ -C ₃ ⁰ , catalytic cracking C ₄ -C ₅ , catalytic cracking naphtha, LCO, And heavy cycle oil, where C ₂ ⁼ −C ₃ ⁼ is part of the desired product and catalytic cracking C ₄ -C ₅ is recycled to the catalytic cracking reactor. ;
(3) Hydrotreating C ₂ ⁰ -C ₄ ⁰ and hydrotreating naphtha in step (1) and catalytic cracking C ₂ ⁰ -C ₃ ^{0 in} step (2) are steamed at a temperature of 700 ° C. to 1000 ° C. And the reaction product vapor is separated, and H ₂ , CH ₄ , steam cracking C ₂ ⁼ -C ₃ ⁼ , steam cracking C ₂ ⁰ -C ₃ ⁰ , steam cracking C ₄ -C ₅ , steam cracking naphtha, and obtain a fuel oil, wherein the steam cracking C ₂ ⁼ -C ₃ ⁼ is part of the desired product, and the steam cracking C ₄ -C ₅ is recycled to the catalytic cracking reactor ;
(4) The catalytic cracking naphtha of step (2) and the steam cracking naphtha of step (3) are first fed to selective hydrogenation and then to solvent extraction to obtain aromatic hydrocarbons and extracted raffinate, The aromatic hydrocarbon is part of the desired product and the extracted raffinate returns to step (3) as one of the raw materials for steam cracking.

The apparatus provided by the present invention includes:
(1) hydrotreating unit, wherein, after the raw material is contacted with hydrogen and a hydrotreating _catalyst, H 2, CH _4, hydrotreated _C ^{2 0} _-C ⁴ 0, hydrotreated naphtha, and hydrotreated A product is obtained;
(2) catalytic cracking unit, wherein the hydroprocessing product is contacted with a catalytic cracking catalyst. The reaction is carried out in at least two reaction zones, and the reaction temperature of at least one of the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone, and WHSV is lower than that of the first reaction zone. The spent catalyst is separated from the reaction product vapor, where after the catalyst is stripped and regenerated, it returns to all or part of the reaction zone of step (2) and the reaction product vapor is separated. Thus, H ₂ , CH ₄ , catalytic cracking C ₂ ⁼ -C ₃ ⁼ , catalytic cracking C ₂ ⁰ -C ₃ ⁰ , catalytic cracking C ₄ -C ₅ , catalytic cracking naphtha, LCO, and HCO are obtained. The _C ^{2 =} _-C ³ = is part of the desired product, and the _C 4 -C ₅ is recycled to the catalytic cracking reactor;
(3) Steam cracking unit, wherein the hydrotreated C ₂ ⁰ -C ₄ ⁰ , hydrotreated naphtha, and catalytic cracking C ₂ ⁰ -C ₃ ⁰ are steamed at a temperature of 700 ° C to 1000 ° C. contact, and the reaction product vapor is _separated, H 2, CH _4, steam cracking _C ^{2 =} _-C ³ =, steam cracking _C ^{2 0} _-C ³ 0, steam cracking _C 4 -C _5, steam cracking naphtha and A fuel oil is obtained, where the steam cracking C ₂ ⁼ -C ₃ ⁼ is part of the desired product;
(4) a selective hydrogenation unit, wherein the catalytic cracking naphtha and steam cracking naphtha are first fed to selective hydrogenation to obtain a selectively hydrotreated naphtha;
(5) Solvent extraction unit, where the selectively hydrotreated naphtha is fed to solvent extraction to obtain aromatic hydrocarbons and extracted raffinate. The aromatic hydrocarbon is part of the desired product and the extracted raffinate returns to the steam cracking unit as one of the raw materials for steam cracking.

Detailed Description of the Invention The method of the present invention is achieved in particular in the following manner:
Embodiment 1
The process for producing light olefins and aromatic hydrocarbons provided by the present invention includes:
The feedstock contacts the catalytic cracking catalyst and reacts under conditions of reaction temperature of 400 ° C. to 800 ° C. and WHSV of 0.1 h ^{−1 to} 750 h ⁻¹ . The reaction is carried out in at least two reaction zones, and the reaction temperature of at least one of the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone, and WHSV is lower than that of the first reaction zone. The spent catalyst is separated from the reaction product vapor and the catalyst is regenerated and then returned to the reactor. The reaction product vapor is separated to yield the desired product, light olefins and aromatic hydrocarbons.

  The process is described in detail as follows.

1) Feedstock The feedstock in each reaction zone is petroleum hydrocarbons and / or other mineral oils, where the petroleum hydrocarbons are vacuum gas oil (VGO), atmospheric gas oil (atmospheric gas oil) (AGO), coked gas oil (CGO), deasphalted oil (DAO), vacuum residuum (VR), atmospheric residuum (AR), cycle oil , Slurry, diesel, naphtha, hydrocarbons having 4 to 8 carbon atoms, alkanes having 2 to 3 carbon atoms, or mixtures thereof, and the other mineral oil is coal Liquid product from liquefaction, tar sand oil or shale oil.

  A preferred feedstock for the first reaction zone is selected from the group consisting of vacuum gas oil, atmospheric gas oil, coke gas oil, deasphalted oil, vacuum residue, atmospheric residue, cycle oil, slurry, diesel, naphtha or mixtures thereof. It is one.

  The preferred feed of the reaction zone downstream of the first reaction zone is from cycle oil, slurry, diesel, naphtha, hydrocarbons having 4-8 carbon atoms, alkanes having 2-3 carbon atoms or mixtures thereof. Is selected from the group consisting of

  Here, VGO, AGO, CGO, DAO, VR, AR, diesel and naphtha are either non-hydrotreated or hydrotreated whole or partial fractions.

  The naphtha is one selected from catalytic cracking naphtha, catalytic cracking naphtha, straight run naphtha, coked naphtha, steam cracking naphtha, pyrolysis naphtha and hydrotreated naphtha, or mixtures thereof, wherein The catalytic cracking naphtha may be derived from the catalytic cracking process of the present invention or may be derived from conventional catalytic cracking, and straight-run naphtha, coke naphtha, steam cracking naphtha, pyrolysis naphtha and hydrotreated naphtha are: Derived from outside this process.

  The diesel is selected from catalytic cracking LCO, straight-run diesel, coked diesel, pyrolysis diesel, and hydrotreated diesel obtained by the process, or mixtures thereof Where the catalytic cracking LCO may be from the catalytic cracking process of the present invention or from conventional catalytic cracking, straight run diesel, coke diesel, pyrolysis diesel, and hydrogenation Treated diesel comes from outside this process.

  Hydrocarbons having 4 to 8 carbon atoms and alkanes having 2 to 3 carbon atoms can be derived from the catalytic cracking process of the present invention or from processes such as conventional catalytic cracking, coking, pyrolysis, hydrogenation, etc. .

  The raw materials for each reaction region may be the same or different. Raw materials for the first reaction zone include heavier hydrocarbons such as VGO, AGO, CGO, DAO, VR, AR, self-produced cycle oil, and self-produced slurry. , Outside cycle oil, outside slurry, diesel and naphtha are preferred. For reaction zones with higher reaction temperatures, lighter hydrocarbons are preferred, such as hydrocarbons having 4 to 8 carbon atoms, alkanes having 2 to 3 carbon atoms, naphtha and diesel.

2) Catalytic cracking catalyst The catalytic cracking catalyst comprises zeolite, inorganic oxide and optional clay, which are each in the following percentage amounts of the total weight of the catalyst: 10% to 50% by weight of zeolite , 5% to 90% by weight of inorganic oxide, and 0% to 70% by weight of clay.

  Here, the zeolite is an active ingredient selected from an intermediate pore size zeolite and optionally a large pore zeolite. The intermediate pore size zeolite accounts for 25% to 100% of the total weight of the zeolite, preferably 50% to 100%, and the large pore zeolite is 0% to 75% of the total weight of the zeolite, preferably Occupies 0% to 50%. The intermediate pore size zeolite is selected from ZSM series zeolites and / or ZRP zeolites, ie ZSM and ZRP zeolites modified with non-metallic elements (eg phosphorus) and / or transition metal elements (eg iron, cobalt and nickel) The See US Pat. No. 5,232,675 for a more detailed description of ZRP zeolite. ZSM series zeolites consist of ZSM-5, ZSM-11, ZSM-12, ZSM-23, ZSM-35, ZSM-38, and ZSM-48, and other zeolites with similar structures, or mixtures thereof It is selected from the group. See US Pat. No. 3,702,886 for a more detailed description of ZSM-5 zeolite. The large pore zeolite is selected from the group consisting of rare earth Y (REY), rare earth HY (REHY), ultrastable Y (ultrastable Y) and high silica Y obtained by various processes, or mixtures thereof. It is.

The inorganic oxide as the binder is selected from silica (SiO ₂ ) and / or alumina (Al ₂ O ₃ ).

  The clay as a matrix (ie supporter) is selected from kaolin clay and / or halloysite.

  The catalytic cracking catalyst in each reactor may be the same or different.

3) Catalytic cracking reactor The cracking reactor comprises at least 2, preferably 2-7, more preferably 2-3 reaction zones, and each reaction zone comprises a riser, fluidized bed, upward transport line (ascending transfer line), descending transfer line and moving floor or parts thereof. The connection method between the various reaction zones is series, parallel or continuous-parallel. The riser can be a conventional unidiameter or various types of taper risers. The structure and dimensions of each reaction zone may be the same or different, where the gas velocity in the fluidized bed is between 0.1 m / s and 2.4 m / s (catalyst is ignored) ).

The first reaction region is a region where a raw material having a boiling point range of 25 ° C. to 1200 ° C. contacts the catalyst. The feedstock for the first reaction zone may also be a conventional catalytic cracking feedstock, where a small amount of light hydrocarbons (eg, C ₂₎ having a lower boiling point than that of the conventional catalytic cracking feedstock. ~C ₆₎ may be included. There may be other reaction regions upstream of this reaction region. The raw materials for the conventional catalytic cracking are vacuum gas oil, atmospheric gas oil, coked gas oil, deasphalted oil, vacuum residuum, atmospheric residue (Atmospheric residuum), cycle oil, slurry, diesel, naphtha, or mixtures thereof.

  The first reaction zone is preferably a riser, upper transport line, lower transport line, or fluidized bed, and more preferably a riser or fluidized bed; the other reaction zone is preferably a fluidized bed, riser And moving beds, more preferably fluidized beds or risers.

4) operating conditions operating conditions are as follows: reaction temperature of 400 ° C. to 800 ° C., preferably from ^{500 ℃ ~700 ℃, WHSV 0.1h -1} ~750h -1, preferably ^{1h -1 ~500h} -1 The reaction pressure is 0.10 MPa to 1.0 MPa (absolute pressure), and the weight ratio of the catalytic cracking catalyst to the raw material is 1 to 150. According to the common knowledge of those skilled in the art, the reaction temperature of a tube reactor (eg riser reactor) means the outlet temperature, and the reaction temperature of a bed reactor (eg fluidized bed reactor) means the average temperature of the bed .

  To perform the fluidization operation, a lifting medium can be injected from the bottom of the reaction zone, and the lifting medium is selected from steam or dry gas, and steam is preferred. The ratio of steam to the raw material is 0.05 to 1.0 by weight.

  The reaction temperature of at least one of the reaction zones downstream of the first reaction zone, preferably 1-6, and more preferably 1-2 is higher than that of the first reaction zone. The difference between the reaction temperature of the reaction zone having a higher reaction temperature and that of the first reaction zone is 10 ° C to 200 ° C, preferably 20 ° C to 100 ° C.

  In order for the reaction temperature of at least one of the reaction zones downstream of the first reaction zone to be higher than that of the first reaction zone, a hot regenerated catalyst, Choose from supplying hot coke deposited catalyst, hot fresh catalyst, and hot feedstock, or having heating coil pipes in the reaction zone Heat is replenished to the reaction zone by one or more means in the manner to be performed. The total amount of heat to be replenished accounts for 10% to 80%, preferably 20% to 60%, of the total reaction system reaction heat.

  The WHSV of at least one of the reaction zones downstream of the first reaction zone, preferably 1-6, and more preferably 1-2 is lower than that of the first reaction zone, and the reaction The ratio of the area WHSV to that of the first reaction area is from 1: 1.1 to 1: 750, preferably from 1: 1.1 to 1: 300.

  The spent catalyst is separated from the reaction product vapor by conventional disengaging, cyclone separation, and after the catalyst is regenerated by burning coke (before this, The catalyst is fed to stripping as needed) and returned to the reactor.

5) Separation of the product The light olefin is ethylene, propylene and optional butene, ie the light olefin is ethylene and propylene or ethylene, propylene and butene.

The method for separating ethylene from the reaction product vapor is the same as that for separating ethylene well known to those skilled in the art, and the method for separating propylene and optional butene from the reaction product vapor is , Well known to those skilled in the art, identical to that for separating propylene and optional butene. The process for separating aromatic hydrocarbons from the cracked naphtha fraction of the reaction product vapor is the same as that for separating aromatic hydrocarbons (ie, extraction) well known to those skilled in the art. Before the separation of aromatic hydrocarbons from a catalytic cracking naphtha obtained by the present process, C ₅ -C ₈ of naphtha can be first isolated as a recycle stream.

The preferred technical scheme of this embodiment includes the following steps:
(1) Raw material and steam enter the first reaction zone (ie riser), contact with regenerated and / or fresh catalytic cracking catalyst, and reaction temperature 500 ° C. to 700 ° C., WHSV 0.1 h ^{−1 to} 750 h ⁻¹ Reaction under conditions of a reaction pressure of 0.1 Mpa to 1.0 Mpa (absolute pressure), a weight ratio of catalytic cracking catalyst to raw material of 1 to 150, and a weight ratio of steam to raw material of 0.05 to 1.0;
(2) Reaction effluent from the first reaction zone enters the second reaction zone (ie, fluidized bed) and contacts the regenerated catalytic cracking catalyst, steam, ethane, propane, C ₄ -C ₈ And a reaction temperature of 500 ° C. to 700 ° C., WHSV 0.1 h ^{−1 to} 750 h ⁻¹ , reaction pressure of 0.1 Mpa to 1.0 Mpa (absolute pressure), weight ratio of catalytic cracking catalyst to raw material of 1 to 150, and raw material It reacts under the condition of a weight ratio of steam to 0.05 to 1.0. The difference between the reaction temperature of the second reaction zone and that of the first reaction zone is 10 ° C. to 200 ° C., preferably 20 ° C. to 100 ° C., and the ratio of WHSV of the second reaction zone to that of the first reaction zone Is from 1: 1.1 to 1: 750, preferably from 1: 1.1 to 1: 300;
(3) separating spent catalyst in the second reaction zone from reaction product vapor, where the spent catalyst enters the regenerator after stripping and is regenerated by burning the coke. , Return to the first reaction zone and the second reaction zone, and the reaction product vapor is separated to obtain the desired products, light olefins and aromatic hydrocarbons.

A preferred technical scheme may further comprise step (4): recycle stream the gas and liquid remaining in the reaction gas-oil other than the desired product, H ₂ and CH _4. As used, this includes ethane, propane, and C ₄ -C ₆ . The extracted raffinate obtained after solvent extraction of the naphtha returns to the second reaction zone and the diesel, cycle oil, and slurry return to the first reaction zone.

Embodiment 2
Another method for producing light olefins and aromatic hydrocarbons provided by the present invention includes:

(1) The raw material and optional recycle stream first enter the hydrotreating unit and come into contact with the hydrotreating catalyst and hydrogen, and the hydrogen partial pressure of 3.0 MPa to 20.0 MPa, reaction of 300 ° C. to 450 ° C. It reacts under conditions of temperature, hydrogen / oil volume ratio of 300-2000, and LHSV of 0.1 h- ^{1 to} 3.0 h- ¹ . Separating the reaction effluent to obtain the hydroprocessing product and recycling the hydrogen for use;
(2) The hydrotreated product is contacted with a catalytic cracking catalyst and reacted in at least two reaction zones under conditions of a reaction temperature of 400 ° C. to 800 ° C. and a WHSV of 0.1 h ^{−1 to} 750 h ⁻¹ Wherein the reaction temperature of at least one of the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone and is downstream of the first reaction zone. The WHSV of at least one of the reaction zones is lower than that of the first reaction zone. The spent catalyst is separated from the reaction product vapor. The spent catalyst is regenerated and then recycled for use, and the reaction product vapor is separated to yield the desired products, light olefins and aromatic hydrocarbons.

This embodiment may further comprise step (3): recycling the desired products, gases and liquids remaining in the reaction gas-oil other than H ₂ and CH ₄ Used as a stream, where the gas recycle stream is ethane, propane and C ₄ and the liquid recycle stream is C ₅ -C ₆ , naphtha extract raffinate, recycled oil and slurry. Ethane, propane and _C 4 -C ₆ and / or extraction raffinate naphtha, diesel, cycle oil and slurry. Ethane, propane, C 4 _{-C 6,} extraction raffinate described above, diesel or hydrotreated diesel, and the slurry, returns to the reaction zone of step (2), and the cycle oil returns to hydrotreating unit. Lower alkanes that do not participate in the reaction are removed from the unit.

  The feedstock is petroleum hydrocarbon and / or other mineral oil, wherein the petroleum hydrocarbon is selected from the group consisting of VGO, AGO, CGO, DAO, VR, AR, diesel and naphtha or mixtures thereof. And the other mineral oil is a liquid product from coal liquefaction, tar sand oil or shale oil.

  Here, VGO, AGO, CGO, DAO, VR, AR, diesel and naphtha are either non-hydrotreated or hydrotreated whole or partial fractions.

  The naphtha is selected from the group consisting of catalytic cracking naphtha, straight run naphtha, coked naphtha, steam cracking naphtha, pyrolysis naphtha and hydrotreated naphtha, or mixtures thereof, wherein contact The cracked naphtha may be derived from the catalytic cracking process of the present invention or may be derived from conventional catalytic cracking, straight run naphtha, coke naphtha, steam cracking naphtha, pyrolysis naphtha and hydrotreated naphtha Derived from.

  The diesel is selected from catalytic cracking LCO, catalytic cracking diesel, straight-run diesel, coked diesel, pyrolysis diesel, and hydrotreated diesel, or mixtures thereof. Yes, where the catalytic cracking LCO can be from the catalytic cracking process of the present invention or from conventional catalytic cracking, straight run diesel, coke diesel, pyrolysis diesel, and hydrotreating Diesel comes from other than this process.

The raw material and cycle oil in step (1) can both enter the hydrotreating reactor after mixing, reducing investment in the facility. The raw material and cycle oil come into contact with the hydrotreating catalyst and hydrogen, and have a hydrogen partial pressure of 3.0 MPa to 20.0 MPa, a reaction temperature of 300 ° C. to 450 ° C., a hydrogen / oil volume ratio of 300 to 2000, and LHSV 0.1 h ⁻ It reacts under the conditions of ^{1 to} 3.0 h ⁻¹ . The reaction effluent is sequentially fed to high pressure separation, low pressure separation, and product fractionation to obtain a hydrotreated product.

It is preferred to hydrotreat heavy oil and cycle oil to obtain optimum reaction effects, however, the high pressure separation, low pressure separation, and product fractionation systems can be shared. In addition, the two reaction systems may use the same pressure level to share fresh and recycled hydrogen compressors. The conditions for hydrogenation of the raw materials are as follows: hydrogen partial pressure 3.0 MPa to 20.0 MPa, reaction temperature 300 ° C. to 450 ° C., hydrogen / oil volume ratio 300 to 2000, and LHSV 0.1 h ^{−1 to} 3. 0h ^-1 ; and the conditions for the hydrogenation treatment of the cycle oil are as follows: hydrogen partial pressure 3.0 MPa to 20.0 MPa, reaction temperature 300 ° C. to 450 ° C., hydrogen / oil volume ratio 300 to 2000, and LHSV 0. 2h- ^{1 to} 2.0h- ¹ .

  The hydrotreating catalyst used in the hydrotreating unit is a group VIB and / or group VIII non-noble metal catalyst supported on alumina and / or amorphous silica-alumina. The non-noble metal of group VIB is selected from Mo and / or W and that of group VIII is selected from Co and / or Ni. High hydrosaturation and denitrogenation activity but low cracking activity are required for this catalyst, retaining as much long linear paraffins in the feed as possible and producing more propylene in the catalytic cracking process. Preferred catalysts are 0-10% by weight additive, 1-9% by weight of one or more Group VIII metals, 12-39% by weight of one or more Group VIB metals, and the remainder alumina and / or as a support. Consists of amorphous silica-alumina, where the additive is selected from non-metallic elements and metallic elements (eg, fluorine, phosphorus, titanium, etc.).

  Compared to the feed, the hydroprocessing product contains less sulfur, nitrogen and aromatic hydrocarbons, and a higher content of hydrogen, and when used as a feed for a catalytic cracking unit, propylene In order to increase the yield of.

  The catalytic cracking catalyst, catalytic cracking reactor, and catalytic cracking operating conditions of this technical scheme are the same as those of Technical Scheme 2, respectively.

Embodiment 3
A third method of light olefins and aromatic hydrocarbons provided by the present invention includes the following:
(1) The raw material is in contact with hydrogen and a hydrotreating catalyst, and a hydrogen partial pressure of 3.0 MPa to 20.0 MPa, a reaction temperature of 300 ° C. to 450 ° C., a hydrogen / oil volume ratio of 300 to 2000, and react under the conditions of LHSV of 1h ^-1 ~3.0h ^-1, and then separating the reaction effluent _{(reaction effluent), H 2,} CH 4, hydrotreated _C ² 0 _-C ⁴ 0, Obtaining a hydrotreated naphtha and a hydrotreated product, wherein the hydrotreated product, hydrogen is recycled for use;
(2) hydrotreated product of step (1) is catalytic cracking catalyst with the contact, and of 400 ° C. to 800 ° C. in at least two reaction zones the reaction temperature and ^0.1h of ^-1 ~750h -1 WHSV of Reacting under conditions, wherein the reaction temperature of at least one reaction zone downstream of the first reaction zone is higher than that of the first reaction zone and at least one of the downstream of the first reaction zone The WHSV of the reaction zone is lower than that of the first reaction zone; the spent catalyst is separated from the reaction product vapor, and the spent catalyst is stripped and regenerated before step (2) Returning to all or part of the reaction zone, and separating the reaction product vapor, H ₂ , CH ₄ , catalytic cracking C ₂ ⁼ -C ₃ ⁼ , catalytic cracking C ₂ ⁰ -C ₃ ⁰ , catalytic cracking C ₄ -C _5, contact Cracked naphtha, light cycle oil, to obtain a heavy cycle oil, wherein the C ₂ ⁼ -C ₃ ⁼ is a part of the desired products, and the catalytic cracking C ₄ -C ₅ Re to the catalytic pyrolysis reactor Circulated;
(3) a catalytic cracking _C ² 0 _-C ^{3 0} hydrotreatment _C ² 0 _-C ^{4 0} and step (2) of step (1) is contacted with steam at a temperature of 700 ° C. to 1000 ° C., and separating the reaction product _vapor, H 2, CH _4, steam cracking _C ^{2 =} _-C ³ =, steam cracking _C ^{2 0} _-C ³ 0, steam cracking _C 4 -C _5, steam cracking naphtha, and fuel oil Where the steam cracking C ₂ ⁼ -C ₃ ⁼ is part of the desired product and the steam cracking C ₄ -C ₅ is recycled to the catalytic cracking reactor;
(4) The catalytic cracking naphtha of step (2) and the steam cracking naphtha of step (3) are first fed to selective hydrogenation and then to solvent extraction to obtain aromatic hydrocarbons and extracted raffinate, The aromatic hydrocarbon is part of the desired product and the extracted raffinate returns to step (3) as one of the raw materials for steam cracking.

  The raw materials, catalytic cracking catalyst, catalytic cracking reactor, and catalytic cracking operation conditions of the technical scheme are the same as those of the technical scheme 2, respectively.

  The solvent used in the solvent extraction is one selected from the group consisting of sulfolane, N-methylpyrrolidone, diethyl glycol ether, triethyl glycol ether, tetraethyl glycol, dimethyl sulfoxide, and N-formylmorpholine and mixtures thereof. The temperature in solvent extraction is 80-120 ° C. and the volume ratio of the solvent to the raw material to be extracted is 2-6. Oil extracted by solvent extraction is one of the desired products, aromatic hydrocarbons, and extracted raffinate (ie, non-aromatic hydrocarbons) is one of the raw materials for steam cracking.

Raw materials for steam cracking are hydrotreated C ₂ ⁰ -C ₄ ⁰ and hydrotreated naphtha in step (1) and catalytic cracking C ₂ ⁰ -C ₃ ⁰ in step (2).

  The reaction conditions in steam cracking are a temperature of 700 ° C. to 1000 ° C., a residence time of 0.05 seconds to 0.6 seconds, and a steam / oil weight ratio of 0.1 to 1.0.

The reaction product vapor is separated, and H ₂ , CH ₄ , steam cracking C ₂ ⁼ -C ₃ ⁼ , steam cracking C ₂ ⁰ -C ₃ ⁰ , steam cracking C ₄ -C ₅ , steam cracking naphtha, fuel oil Obtained, where the steam cracking C ₂ ⁼ -C ₃ ⁼ is part of the desired product. The steam cracking C ₂ ⁰ -C ₃ ⁰ is recycled to the steam cracking reactor, and the steam cracked C ₄ -C ₅ is recycled to the catalytic pyrolysis reactor.

  This technical scheme produces light olefins (eg, propylene, ethylene, etc.) from heavy feeds by integrating hydroprocessing, catalytic cracking, steam cracking, solvent extraction, etc., where propylene Yields over 30% and simultaneously produces aromatic hydrocarbons (eg, toluene, xylene, etc.).

The apparatus provided by the present invention includes:
(1) hydrotreating unit, wherein, after the raw material is contacted with hydrogen and a hydrotreating _catalyst, H 2, CH _4, hydrotreated _C ^{2 0} _-C ⁴ 0, hydrotreated naphtha, and hydrotreated A product is obtained;
(2) catalytic cracking unit, wherein the hydroprocessing product is contacted with a catalytic cracking catalyst and in at least two reaction zones a reaction temperature of 400 ° C. to 800 ° C. and 0.1 h ^{−1 to} 750 h ⁻¹ In which the reaction temperature of at least one reaction zone downstream of the first reaction zone is higher than that of the first reaction zone and downstream of the first reaction zone The WHSV of at least one reaction zone is lower than that of the first reaction zone; the spent catalyst is separated from the reaction product vapor, and the spent catalyst is regenerated by stripping, then the process ( returns to all or a portion of the reaction zone 2), and separates the reaction product _vapor, H 2, CH _4, catalytic cracking _C ² = _-C ^{3 =,} catalytic cracking _C ² 0 _-C ^{3 0} , Catalytic cracking C ₄ -C _5, give catalytic cracking naphtha, light cycle oil, heavy cycle oil, wherein the _C ^{2 =} _-C ³ = is part of the desired product, and the _C 4 -C ₅ Said catalytic cracking Recycled to the reactor;
(3) Steam cracking unit, where the hydrotreated C ₂ ⁰ -C ₄ ⁰ , hydrotreated naphtha, catalytic cracking C ₂ ⁰ -C ₃ ⁰ , and catalytic cracking C ₂ ⁰ -C ₃ ⁰ are 700 ° C. in contact with steam at a temperature of to 1000 ° C., and the reaction product vapor is _separated, H _2, CH _4, steam cracking _C ^{2 =} _-C ³ =, steam cracking _C ^{2 0} _-C ³ 0, steam cracking C ₄ -C _5, to give the steam cracking naphtha, and fuel oil, wherein the steam cracking _C ^{2 =} _-C ³ = is part of the desired product;
(4) a selective hydrogenation unit, wherein the catalytic cracking naphtha and steam cracking naphtha are fed to selective hydrogenation to obtain a selectively hydrotreated naphtha;
(5) A solvent extraction unit, wherein the selectively hydrotreated naphtha is fed to solvent extraction to obtain an aromatic hydrocarbon and an extracted raffinate, where the aromatic hydrocarbon is desired And the extracted raffinate returns to step (3) as one of the raw materials for steam cracking.

  The method of the present invention is further described below in combination with the drawings.

  FIG. 1 is a schematic diagram of the catalytic cracking process of Technical Scheme 1.

  There are 7 reaction zones in this figure, where the temperature of both reaction zone III and reaction zone V is higher than that of reaction zone I and their WHSV is lower than that of reaction zone I . A medium PS2, such as propane, is injected from the catalyst pre-lifting section of the catalyst P6 at the bottom of the reaction zone I and the reaction takes place. This pre-lifting section P6 can also be regarded as a reaction zone.

The process flow is as follows:
The regenerated catalyst C1 from the regenerator enters the catalyst pre-lifting section P6 at the bottom of the reaction zone I through the catalyst pipeline P1, is lifted by the medium PS1, and then reacts with the medium PS2. In reaction zone I, feed S1 is in contact with the stream from pre-lifting section P6 and the catalyst and reaction temperature of 500 ° C. to 650 ° C., WHSV of 2h ^{−1 to} 300h ⁻¹ , catalyst / oil of 3 to 20 It reacts under a weight ratio and a reaction pressure of 0.12 MPa to 0.6 MPa (absolute pressure). The reaction product and catalyst then both enter reaction zone II.

In the reaction zone II, _{C 4} hydrocarbon feedstock S2 from another catalytic cracking unit, in contact with the catalyst and stream from reaction zone 1 and 490 ° C. to 640 ° C. reaction ^temperature, of 20h ^-1 ~750h -1 It reacts under WHSV, a catalyst / oil weight ratio of 3-20, and a reaction pressure of 0.12 MPa to 0.6 MPa (absolute pressure). The reaction product and catalyst then both enter reaction zone III.

In reaction zone III, the C ₄ hydrocarbon feed S3 from the process is in contact with the catalyst and stream from reaction zone II and the regenerated catalyst from pipeline P2, and the reaction temperature of 510 ° C. to 660 ° C., 2h ^−1. It reacts under WHSV of ˜150 h ⁻¹ , a catalyst / oil weight ratio of 3-20 and a reaction pressure of 0.12 MPa to 0.6 MPa (absolute pressure). The reaction product and catalyst then both enter reaction zone IV.

In a reaction zone IV, _{C 5} hydrocarbon feedstock S4 from the process, in contact with the catalyst and stream from reaction zone III, and 490 ° C. to 640 ° C. reaction ^temperature, WHSV, 3 of 20h ^-1 ~750h -1 It reacts under a catalyst / oil weight ratio of ˜20 and a reaction pressure of 0.12 MPa to 0.6 MPa (absolute pressure). The reaction product and catalyst then both enter reaction zone V.

In a reaction zone V, C ₅ hydrocarbon feedstock S5 from the part of the raw material and the process includes a catalyst and stream from reaction zone IV, the spent catalyst C5 from pipeline P5, the regenerated catalyst from pipeline P4 and C4, and the regenerated catalyst C3 from the pipeline P3 is later contacted and reacted with propane PS3 catalyst and the reaction product in contact and mixed, and 510 ° C. to 700 ° C. reaction ^temperature, of 1h ^-1 ~100h -1 It reacts under WHSV, a catalyst / oil weight ratio of 3-50, and a reaction pressure (absolute pressure) of 0.12 MPa to 0.6 MP. The reaction product and catalyst then both enter reaction zone VI.

In a reaction zone VI, the _{C 6} material S6 in the present process, contact and mixed with the catalyst and stream from reaction zone V, and 510 ° C. to 700 ° C. reaction ^temperature, WHSV, 3 of 20h ^-1 ~700h -1 The reaction is carried out under a catalyst / oil weight ratio of ˜50 and a reaction pressure (absolute pressure) of 0.12 MPa to 0.6 MPa. The reaction product and catalyst then both enter reaction zone VII.

In a reaction zone VII, water or steam S7 is, reaction in contact and mixed with the catalyst and stream from the region VI, and 450 ° C. to 700 ° C. reaction ^temperature, WHSV of 20h ^-1 ~700h -1, 3~50 catalyst / Oil weight ratio, and reaction under 0.12 MPa to 0.6 MPa reaction pressure (absolute pressure). The reaction product vapor and catalyst then enter a catalyst / product vapor separation system. The separated product vapor enters the product vapor separation system, and the catalyst is stripped as necessary before entering the regenerator for production.

  FIG. 2 is a schematic diagram of the preferred catalytic cracking process of Technical Scheme 1.

  In this process, the catalytic cracking reactor 40 comprises a first reaction zone (ie, riser A) and a second reaction zone (ie, fluidized bed B) in succession. The difference between the reaction temperature of the second reaction zone and that of the first reaction zone is 10 to 200 ° C., preferably 20 ° C. to 100 ° C., and the ratio of WHSV of the second reaction zone to that of the first reaction zone is 1: 1.1 to 1: 750, preferably 1: 1.1 to 1: 300.

The process flow is as follows:
Pre-lifting steam enters from the bottom of riser A through pipeline 57, and the regenerated catalyst moves upward along riser A while increasing in speed under the pre-lifting action of the steam. Moving. The raw material is injected into the riser A through the pipeline 5 while spraying steam from the pipeline 39, and comes into contact with the regenerated catalyst. After being lifted by steam from pipeline 55, the regenerated catalyst from pipeline 54 enters a vertical transfer line 56 along with C ₂ ⁰ -C ₃ ⁰ and C ₄ -C ₈ from pipeline 26. Enters and travels up and finally enters fluidized bed B where the reaction from the catalyst and product vapor from riser A takes place. Steam enters the bottom of fluidized bed B through pipeline 41 to ensure fluidization and reaction of fluidized bed B. The product vapor produced in fluidized bed B and the deactivated spent catalyst enter a cyclone in a disengager 43 through pipeline 42, where the spent catalyst from the product vapor is Separation takes place. The product vapor enters the recovery chamber 44 and the catalyst fine powder returns to the disengager through the leg. Spent catalyst in the disengager flows to the stripping section 47 and contacts the steam from the pipeline 48. Product vapor stripped from the spent catalyst enters the chamber 44 through a cyclone. The spent catalyst after stripping enters the regenerator 50 through a sloped pipe 49. Main air enters the regenerator, the coke on the spent catalyst is burned, the deactivated spent catalyst is regenerated, and the flue gas is routed to the fume machine via pipeline 52. enter. The regenerated catalyst is divided into two parts, where one part enters the riser A through the inclined pipe 53 and the other part sequentially connects the inclined pipe 54 and the vertical transport line 56 for recycling use. And enters fluidized bed B. Product vapor in chamber 44 enters the next separation system 46, from which ethylene and propylene obtained by separation are recovered via pipeline 7, and catalytic cracking dry gas (ie hydrogen and methane) is piped. Recovered via line 6, catalytic cracking C ₂ ⁰ -C ₃ ⁰ is recovered via pipeline 8 and finally introduced into vertical transport line 26, catalytic cracking C ₄ -C ₅ passing through pipeline 9 And the catalytic cracking naphtha is recovered via the pipeline 10 to separate aromatic hydrocarbons (for example, toluene, xylene, etc.), the catalytic cracking diesel is recovered via the pipeline 11 and the catalytic catalytic cycle oil. And slurry is recovered via pipeline 12. Catalytic cracking C ₄ -C ₅ returns to fluidized bed B sequentially through pipeline 26 and vertical transport line 56. Diesel or hydrotreated diesel, catalytic cracking cycle oil, and slurry together return to riser A through pipeline 5.

  FIG. 3 is a schematic principle flowsheet of the process of Technical Scheme 2.

The principle flow of the process after hydroprocessing is briefly described as follows: The feedstock is mixed with cycle oil from pipeline 15 through pipeline 1 and then hydrogenated through pipeline 22. Entering treatment unit 23, in contact with hydrotreating catalyst and hydrogen, and hydrogen partial pressure of 3.0 MPa to 20.0 MPa, reaction temperature of 300 ° C. to 450 ° C., hydrogen / oil volume ratio of 300 to 2000, and 0 Reacts under conditions of LHSV from ¹ h ^{−1 to} 3.0 h ⁻¹ . The oil produced in the hydroprocessing unit mixes with the recycle stream from the pipeline 17 through the pipeline 24, then enters the catalytic cracking reactor 2, contacts the catalytic cracking and steam, and 500 ° C to 700 ° C. Temperature, 0.15 to 0.4 MPa pressure (absolute pressure), 5 to 50 weight ratio of catalytic cracking catalyst to catalytic cracking raw material, 0.05 to 1.0 weight ratio of steam to catalytic cracking raw material Reacts below. The coked spent catalyst and reaction product vapor enter the catalyst / oil separator 4 through the pipeline 3, and the separated spent catalyst enters the regenerator 6 through the pipeline 5. The catalyst regenerated by burning the coke has a higher activity and selectivity, which returns to the reactor 2 via the pipeline 7 and the reaction product vapor enters the product separator 9. The separated ethylene and propylene are recovered via pipeline 12 and cracked naphtha (from which C ₅ -C ₆ is taken) enters solvent extraction unit 19. The resulting aromatic hydrocarbons are withdrawn through pipeline 21, hydrogen and methane are withdrawn through pipeline 10, ethane and propane are withdrawn through pipeline 11, C ₄ -C ₆ pipe Recovered via line 13, cycle oil is recovered via pipeline 15, and slurry is recovered via pipeline 16. Ethane and propane, C ₄ -C ₆ , naphtha extract raffinate, and slurry are all or partly returned to catalytic cracking reactor 2 sequentially through pipelines 17 and 18, and cycle oil is supplied to pipeline 15 and The process returns to the hydroprocessing unit 23 through 22 sequentially.

  FIG. 4 is a schematic principle flowsheet of the process and apparatus of Technical Scheme 3.

The principle flow of this process is as follows: The feedstock mixes with diesel from pipeline 11 through pipeline 1 and then enters hydrogenation unit A. H ₂ , CH ₄ , hydrotreated C ₂ ⁰ -C ₄ ⁰ , hydrotreated naphtha, and hydrotreated product obtained by hydrotreating are routed through pipelines 2, 3, 4 and 5, respectively. Collected. Hydrotreated C ₂ ⁰ -C ₄ ⁰ and hydrotreated naphtha mix through pipelines 3 and 4, respectively, and then sequentially enter pipelines 23 and 24 into steam cracking unit E. The hydroprocessing product as feedstock for catalytic cracking is mixed with C ₄ -C ₅ from pipeline 26 and then enters catalytic cracking unit B and is at a temperature of 500 ° C.-700 ° C., 0.15- The reaction is carried out under conditions of a pressure of 0.4 MPa (absolute pressure), a weight ratio of the catalytic cracking catalyst to the catalytic cracking raw material of 5 to 50, and a weight ratio of steam to the catalytic cracking raw material of 0.05 to 1.0. _H 2, CH ₄ obtained in catalytic cracking unit B, catalytic cracking _C ^{2 =} _-C ³ =, catalytic cracking _C ^{2 0} _-C ³ 0, catalytic cracking _C 4 -C _5, catalytic cracking naphtha, diesel, a cycle oil, Recover from pipelines 2, 3, 4 and 5, respectively. C ₂ ⁼ -C ₃ ⁼ is one of the desired product. C ₂ ⁰ -C ₃ ⁰ as one of the steam cracking raw materials enters the steam cracking unit E sequentially through the pipelines 8, 22 and 24. Catalytic cracking _C 4 -C ₅ passes through the pipeline 9, and mixed with steam cracking _C 4 -C ₅ from the pipeline 18, then returns to the catalytic cracking unit B sequentially through pipelines 26 and 27. The catalytic cracking naphtha mixes with the steam cracking naphtha from the pipeline 19 and then enters the selective hydrogenation unit C through the pipeline 25. Diesel sequentially returns to hydrogenation unit A through pipelines 11 and 28. The cycle oil returns to the catalytic cracking unit B through the pipeline 12. The stream from the selective hydrogenation unit C enters the solvent extraction unit D. The BTX obtained in the solvent extraction unit D is recovered via the pipeline 14 as the desired product, and the extracted raffinate enters the steam cracking unit E sequentially through the pipelines 13, 22 and 24.

Hydrotreated C ₂ ⁰ -C ₄ ⁰ , hydrotreated naphtha, catalytic cracking C ₂ ⁰ -C ₃ ⁰ , and extracted raffinate are first mixed through pipelines 3, 4, 8 and 13, respectively, and Next, the steam cracking unit E is entered through the pipeline 24. H ₂ , CH ₄ , steam cracking C ₂ ⁼ -C ₃ ⁼ , steam cracking C ₂ ⁰ -C ₃ ⁰ , steam cracking C ₄ -C ₅ , steam cracking naphtha, and fuel oil are pipelines 15, 16, respectively. 17, 18, 19 and 20, where steam cracking C ₂ ⁼ −C ₃ ⁼ is recovered from the unit via pipeline 16 as one of the desired products, and steam cracking C ₂ ⁰ -C ₃ ⁰ returns to steam cracking unit E through pipeline 17, steam cracking C ₄ -C ₅ enters catalytic cracking unit B through pipelines 18, 26 and 27 in sequence, and steam Cracking naphtha passes through pipeline 19 to selective hydrogenation unit C That.

  By using several processes provided by the present invention, refineries can produce light olefins such as propylene, ethylene, etc. up to a maximum, where the yield of propylene is Aromatic hydrocarbons greater than 20% by weight, preferably greater than 25%, and more preferably greater than 30%, and also rich in joint products, toluene, xylene, etc. can be produced. Therefore, refinery concept technical breakthrough, refinery conversion from traditional fuel and fuel-lubricant production mode to chemical industry mode, and single oil refining to the production of high value added downstream products The refinery's development and expansion will be realized. This conversion not only solves the problem of a lack of chemical feedstock, but also increases the benefits of the refinery.

  The following examples further illustrate the invention but do not limit it.

  The raw material used in the examples is VGO, the characteristics are shown in Table 1, and the solvent used in the examples is sulfolane.

The process for producing the catalytic cracking catalyst used in the examples is briefly described as follows:
1) 20 g NH ₄ Cl dissolved in 1000 g water, into which 100 g (dry basis) crystallized product ZRP-1 zeolite (produced at the Catalyst Plant of Qil Petrochemicals Co., SiO ₂ / Al ₂ O ₃ = 30, content of rare earth RE ₂ O ₃ = 4.0 wt%) and after 0.5 hour exchange reaction and filtration at 90 ° C., a filter cake was obtained. The filter cake was impregnated with a solution of 4.0 g H ₃ PO ₄ (85% concentration) and 4.5 g Fe (NO ₃ ) ₃ in 90 g water and then dried. The obtained solid was calcined at 550 ° C. for 2 hours to obtain an intermediate pore size zeolite containing phosphorus and iron and having a structure of MFI. The composition determined by elemental analysis _{was 0.1Na 2 O · 5.1Al 2 O 3} · 2.4P 2 O 5 · 1.5Fe 2 O 3 · 3.8RE 2 O 3 · 88.1SiO 2 .

2) 75.4 kg of halloysite (product of Suzuhou Porcelain Clay Co., 71.6 wt% solids) was slurried with 250 kg of decationized water and 54.8 kg of pseudo-boehmite (industrial product of Shandong Alumina Plant) , 63 wt% solids). The pH was adjusted to 2-4 with hydrochloric acid, the slurry was stirred uniformly, and allowed to stand for aging at 60-70 ° C. for 1 hour. The temperature was lowered below 60 ° C. while maintaining the pH at 2-4, then 4.5 kg of alumina sol (Shandong Alumina Plant industrial product, Al ₂ O ₃ content 21.7 wt%) was added. After stirring for 40 minutes, a mixed slurry was obtained.

3) Intermediate pore size zeolite (45 kg dry basis) containing phosphorus and iron and having MFI structure prepared in step 1) and DASY zeolite (the Catalyst Plant of Qilu Petrochemicals Co. industrial product, unit cell size 2.445) ˜2.448 nm, RE ₂ O ₃ content 2.0%, dry base 7.5 kg) was added to the mixed slurry obtained in step 2) and stirred uniformly. The resulting slurry was shaped by spray drying and the product was washed with a solution of ammonium dihydrogen phosphate (phosphorus content 1% by weight) to remove free Na ⁺ . After drying, a sample of catalytic cracking catalyst was obtained. The composition of the catalyst was 30% by weight MFI structure intermediate pore size zeolite containing phosphorus and iron, 5% by weight DASY zeolite, 23% by weight pseudo-boehmite, 6% by weight alumina sol, and the remaining kaolin.

The process for preparing the hydrotreating catalyst used in the examples is briefly described as follows: ammonium metatungstate (NH ₄ W ₄ O ₁₃ · 18H ₂ O, chemically pure) and Nickel nitrate (Ni (NO ₃ ) ₂ .6H ₂ O, chemically pure) was dissolved in water to prepare a 200 ml solution. 50 g of alumina support was added to the solution and impregnated at room temperature for 3 hours. During impregnation, the impregnation solution was treated with ultrasonic equipment for 30 minutes, then cooled, filtered and dried in a microwave oven for about 15 minutes. The composition of the _{catalyst, WO} 3 30 wt%, NiO 3.1 wt%, and was remaining alumina.

The process for preparing the hydrotreating catalyst used in the examples is briefly described as follows: ammonium metatungstate (NH ₄ W ₄ O ₁₃ · 18H ₂ O, chemically pure) and nickel nitrate _{(Ni (NO 3) 2 ·} 6H 2 O, chemically pure) were dissolved in water to prepare a solution of 200 mL. 100 g of alumina support was added to the solution and impregnated for 4 hours at room temperature. After isolation, the wet catalyst was dried in an oven for 4 hours and calcined in a tube furnace with blowing air at 500 ° C. for 4 hours. The composition of the catalyst was WO ₃ 25.3% by weight, NiO 2.3% by weight, and the remaining alumina.

Example 1
The experiment of this example was performed according to the flow in FIG. Raw material A was used directly as the raw material for catalytic cracking and the experiment was conducted in 7 reaction zones consisting of riser and fluidized bed. Reaction temperatures in reaction zones I, II, III, IV, V, VI and VII are 530 ° C, 520 ° C, 550 ° C, 540 ° C, 640 ° C, 620 ° C and 580 ° C, respectively. There, the reaction region I, II, III, IV, V, VI and VII of WHSV in ^{turn, 360h -1, 720h -1, 20h} -1, 180h -1, 5h -1, 200h -1 and 620H ^{- 1,} where the reaction temperature of the reaction zone III and V, respectively, 20 ° C. and 110 ° C. higher than that of the reaction zone 1, WHSV and reaction zone I (riser reaction zone III and V (fluidized bed) ) Ratio to that of 1:18 and 1:72, respectively, and the heat replenished to reaction zones III and V is 11% and 60% of the total reaction heat, respectively. Occupy. Finally, the product is isolated, wherein the C ₃ -C _5, was circulated to the flowable floor. The operating conditions of reaction zones I, III and V and the product distribution are shown in Table 2.

  From Table 2, the yields of propylene and ethylene reach 35.21 wt% and 14.56 wt%, respectively, and those of toluene and xylene are 3.95 wt% and 4.26 wt%, respectively. It can be understood that there is.

Example 2
The experiment of this example was performed according to the flow in FIG. Raw material B is used directly as the raw material for catalytic cracking and experiments are performed in an intermediate size riser and fluid bed reactor, where the reaction temperature of the fluid bed is 30 ° C. higher than that of the riser, the fluid bed The ratio of WHSV to that of the riser is 1: 360, and the heat supplied to the fluidized bed accounts for 22% of the total reaction heat. Finally, the product was separated where only the slurry was circulated to the riser and C ₄ -C ₆ was circulated to the fluidized bed, but no other streams were circulated. Table 2 shows the catalytic cracking operating conditions and product distribution.

  From Table 2, the yields of propylene and ethylene reach 30.46 wt% and 18.31 wt%, respectively, and those of toluene and xylene are 2.45 wt% and 7.38 wt%, respectively. It can be understood that there is.

Example 3
The experiment of this example was performed according to the flow in FIG. Raw material A is first subjected to hydroprocessing, and the hydroprocessing product (hydrogen content increased from 12.40 wt% to 13.54 wt% and aromatic hydrocarbon content is , 44.1 wt% to 20.0 wt%) was used as a raw material for catalytic cracking. Experiments were performed in an intermediate size riser and fluidized bed reactor, where the reaction temperature of the fluidized bed was 40 ° C. higher than that of the riser and the ratio of WHSV of the fluidized bed to that of the riser was 1:30. And the heat supplied to the fluidized bed accounts for 25% of the total reaction heat. Finally, the product was separated, where only the slurry was circulated to the riser, but no other streams were circulated. The operating conditions and product distribution for hydroprocessing and catalytic cracking are shown in Table 3.

  From Table 3, the yields of propylene and ethylene reach 32.97 wt% and 12.63 wt%, respectively, and those of toluene and xylene are 1.93 wt% and 4.05 wt%, respectively. It can be understood that there is.

Example 4
The experiment of this example was performed according to the flow in FIG. The reactor for catalytic cracking is a fluidized bed reactor and riser in a pilot plant, where the reaction temperature of the fluidized bed is 40 ° C. higher than that of the riser, the WHSV of the fluidized bed and that of the riser. The ratio is 1:30, and the heat supplied to the fluidized bed accounts for 30% of the total reaction heat. Finally, the product was separated, where only the slurry was circulated to the riser, but no other streams were circulated. Hydrotreatment, catalytic cracking, selective hydrogenation and solvent extraction experiments were all performed in the corresponding intermediate size units.

The raw materials used in this example are the same as in Example 3 (ie, raw material A), and the operating conditions and product distribution are shown in Table 3. From Table 3, the yields of propylene and ethylene, respectively, was 40.65% and 20.64% by weight, and they toluene and _{C 8} aromatic hydrocarbons, respectively, 4.34 wt% and 5 It can be seen that it is 18% by weight.

FIG. 1 is a schematic diagram of the catalytic cracking process of Technical Scheme 1. FIG. 2 is a schematic diagram of the preferred catalytic cracking process of Technical Scheme 1. FIG. 3 is a schematic principle flow sheet of the process of Technical Scheme 2. FIG. 4 is a schematic principle flow sheet of the process and apparatus of Technical Scheme 3.

Claims (29)

A process for the production of light olefins and aromatics comprising a catalytic cracking catalyst under conditions of a reaction temperature of 400 ° C. to 800 ° C. and a weight space velocity (WHSV) of 0.1 h ^{−1 to} 750 h ⁻¹ Reacting the feedstock by contacting; separating the spent catalyst from the reaction product vapor; regenerating the spent catalyst; and returning the regenerated catalyst to a reactor, the reaction comprising: Carried out in at least two reaction zones, wherein the reaction temperature of at least one of the reaction zones downstream of the first reaction zone is higher than that of the first reaction zone and its WHSV but lower than that of the first reaction zone; the difference between that of the reaction temperature and the first reaction zone of the reaction region having a high reaction temperature by lined, 10 ° C. to 20 A ° C., and the ratio of the WHSV and the first reaction zone of the reaction zone is 1: 1.1 to 1: characterized in that it is a 750, method. The method according to claim 1, wherein the first reaction zone is a reaction zone where a raw material having a boiling point range of 25 ° C. to 1200 ° C. is in contact with the catalyst. The reaction is performed in 2 to 7 reaction zones, wherein the reaction temperature of 1 to 6 reaction zones downstream of the first reaction zone is higher than that of the first reaction zone. The process according to claim 1, characterized in that they are high and their WHSV is lower than that of the first reaction zone. The reaction is performed in 2 to 3 reaction zones, wherein the reaction temperature of 1 to 2 reaction zones downstream of the first reaction zone is higher than that of the first reaction zone. The process according to claim 1, characterized in that they are high and their WHSV is lower than that of the first reaction zone. The difference between the reaction temperature of the reaction zone having a higher temperature and that of the first reaction zone is 20 ° C. to 100 ° C., and the ratio of WHSV of this reaction zone to that of the first reaction zone is 1 The method according to claim 1, wherein the method is 1.1 to 1: 300. The heat source for the reaction zone having a higher reaction temperature is a hot regenerated catalyst, a hot coke deposited catalyst, a hot fresh catalyst, and a hot feedstock. ), As well as one or more selected from the group consisting of one or more heating coil tubes provided in the reaction zone, or a mixture thereof. The process according to claim 1, characterized in that the total amount of heat supplied to the reaction zone having a higher reaction temperature accounts for 10% to 80% of the reaction heat of the total reaction system. The process according to claim 1, characterized in that the total amount of heat supplied to the reaction zone having a higher reaction temperature accounts for 20% to 60% of the reaction heat of the total reaction system. Conditions in each of the reaction regions are a reaction temperature of 500 ° C. to 700 ° C., WHSV 1h ^{−1 to} 500h ⁻¹ , a reaction pressure of 0.10 MPa to 1.0 MPa, and a weight ratio of the catalytic cracking catalyst to the raw material of 1 to 150. The method according to claim 1, wherein: Each of the reaction zones is selected from a riser, a fluidized bed, an ascending transfer line, a descending transfer line, and a moving bed, and a method for connecting the various reaction zones, 2. A method according to claim 1, characterized in that it is series, parallel or mixed. The feedstock is petroleum hydrocarbon and / or other mineral oil, wherein the petroleum hydrocarbon is a vacuum gas oil, an atmospheric gas oil, a coked gas oil, a deasphalted oil, a vacuum Residual oil (vacuum residuum), atmospheric residuum, cycle oil, slurry, diesel, naphtha, hydrocarbons with 4-8 carbon atoms, alkanes with 2-3 carbon atoms or mixtures thereof The one selected from the group and characterized in that the other mineral oil is a liquid product from coal liquefaction, tar sand oil or shale oil. The method according to 1. The raw material of the first reaction zone is selected from the group consisting of vacuum gas oil, atmospheric gas oil, coke gas oil, deasphalted oil, vacuum residue, atmospheric residue, cycle oil, slurry, diesel, naphtha or mixtures thereof. The method according to claim 1, characterized in that: The reaction zone feed downstream of the first reaction zone is comprised of cycle oil, slurry, diesel, naphtha, hydrocarbons having 4-8 carbon atoms, alkanes having 2-3 carbon atoms, or mixtures thereof. The method according to claim 1, wherein the method is selected from: The vacuum gas oil, atmospheric gas oil, coke gas oil, deasphalted oil, vacuum residue, atmospheric residue, diesel and naphtha are all or a fraction thereof. 12. The method according to 12 . The naphtha is selected from the group consisting of catalytic cracking naphtha, straight naphtha, coked naphtha, steam cracking naphtha, pyrolysis naphtha and hydrotreated naphtha, or mixtures thereof. The method according to claim 11, 12 or 13 . The diesel is selected from the group consisting of catalytic cracking light cycle oil (LCO), straight run diesel, coked diesel, pyrolysis diesel and hydrotreated diesel, or mixtures thereof. The method according to claim 11, 12 or 13 . The catalytic cracking catalyst comprises zeolite, an inorganic oxide and optional clay, and the content of various components in the catalyst is 10 wt% to 50 wt% of zeolite, 5 wt% to 90 wt% of inorganic oxide The method according to claim 1, characterized in that the clay is from 0% to 70% by weight. The zeolite is a medium pore size zeolite and optionally a large pore zeolite, wherein the medium pore size zeolite accounts for 25% to 100% of the total weight of the zeolite, and the large pore zeolite is Accounts for 0% to 75% of the total weight of the zeolite; the intermediate pore size zeolite is selected from ZSM series zeolite and / or ZRP zeolite, and the ZSM series zeolite is ZSM-5, ZSM-11, ZSM-12 , ZSM-23, ZSM-35, ZSM-38, and ZSM-48, or mixtures thereof; the large pore zeolite is rare earth Y, rare earth HY, ultrastable Y (Ultrastable Y), and high silica Y, or a mixture thereof. The method according to claim 17 , wherein: Wherein the inorganic oxide is selected from SiO ₂ and / or Al ₂ O _3, and the clay, characterized in that it is selected from kaolin clay and / or halloysite, The method of claim 17. The process according to claim 1, characterized in that the lower olefin is ethylene, propylene and optional butene. A process for producing light olefins and aromatic hydrocarbons comprising the following steps:
(1) The raw material and optional recycle stream enter the hydrotreating unit, come into contact with the hydrotreating catalyst and hydrogen, and the hydrogen partial pressure of 3.0 MPa to 20.0 MPa, the reaction temperature of 300 ° C. to 450 ° C. , Reaction under conditions of hydrogen / oil volume ratio of 300-2000, and liquid hourly space velocity (LHSV) of 0.1 h ^{−1 to} 3.0 h ⁻¹ , separating reaction effluent and hydrogenating A treated product is obtained and hydrogen is circulated for reuse;
(2) supplying the hydroprocessing product to catalytic cracking according to the method of claim 1;
Including the method. A process for producing light olefins and aromatic hydrocarbons comprising:
(1) The raw material is in contact with hydrogen and a hydrotreating catalyst, and a hydrogen partial pressure of 3.0 MPa to 20.0 MPa, a reaction temperature of 300 ° C. to 450 ° C., a hydrogen / oil volume ratio of 300 to 2000, and Reacts under conditions of liquid hourly space velocity (LHSV) of ¹ h ^{−1 to} 3.0 h ⁻¹ and then the reaction effluent is separated and H ₂ , CH ₄ , hydrotreated C ₂ ⁰ -C 4 ^_0, resulting hydrotreated naphtha, and hydrotreated product, and is circulated for re-use of hydrogen;
(2) Feeding the hydroprocessed product of step (1) to catalytic cracking according to the method of claim 1 wherein the spent catalyst is separated from the reaction product vapor, After being stripped and regenerated, it returns to all or part of the reaction zone of step (2), and the reaction product vapor is separated to yield H ₂ , CH ₄ , catalytic cracking C ₂ ⁼ -C ₃ ⁼ , Catalytic cracking C ₂ ⁰ -C ₃ ⁰ , catalytic cracking C ₄ -C ₅ , catalytic cracking naphtha, LCO, and heavy cycle oil (HCO) are obtained, where the C ₂ ⁼ -C ₃ ⁼ is the desired product And the C ₄ -C ₅ is recycled to the catalytic cracking reactor;
(3) a catalytic cracking _C ² 0 _-C ^{3 0} of the hydrotreated _C ² 0 _-C ^{4 0} and hydrogenated naphtha and step (2) of step (1), steam at a temperature of 700 ° C. to 1000 ° C. And the reaction product vapor is separated, and H ₂ , CH ₄ , steam cracking C ₂ ⁼ -C ₃ ⁼ , steam cracking C ₂ ⁰ -C ₃ ⁰ , steam cracking C ₄ -C ₅ , steam give a cracking naphtha, and fuel oil, wherein the steam cracking C ₂ ⁼ -C ₃ ⁼ is part of the desired product, and the steam cracking C ₄ -C ₅ is circulated to the catalytic cracking reactor ;
(4) The catalytic cracking naphtha of step (2) and the steam cracking naphtha of step (3) are first fed to selective hydrogenation and then to solvent extraction to obtain aromatic hydrocarbons and extracted raffinate, The aromatic hydrocarbon is part of the desired product and the extracted raffinate returns to step (3) as one of the raw materials for steam cracking,
Including the method. The hydrotreating catalyst according to step (1) is alumina and / or amorphous Syria Ca - characterized in that it is a non-noble metal catalysts Group VIB and / or Group VIII supported on alumina, claim 21 Or the method according to 22 . The method according to claim 23 , characterized in that the non-noble metal of group VIB is selected from Mo and / or W and that of group VIII is selected from Co and / or Ni. The conditions for steam cracking in step (3) are: a temperature of 700 ° C. to 1000 ° C., a residence time of 0.05 seconds to 0.6 seconds, and a weight ratio of steam to oil of 0.1 to 1.0. 23. The method of claim 22 , wherein there is. The conditions for selective hydrogenation in step (4) are: hydrogen pressure of 1.2 MPa to 8.0 MPa, reaction temperature of 150 ° C. to 300 ° C., volume ratio of hydrogen to oil of 150 to 600, and 1 h ⁻ The method according to claim 22 , characterized in that the LHSV is between ^{1 and} 20 h ⁻¹ . The solvent described in the step (4) is selected from the group consisting of sulfolane, N-methylpyrrolidone, diethyl glycol ether, triethyl glycol ether, tetraethyl alcohol, dimethyl sulfoxide, N-formylmorpholine, and mixtures thereof. The method according to claim 22 , characterized in that: Step (4) in a temperature 80 ° C. to 120 ° C. for solvent extraction, and wherein the ratio of solvent to feed for the extraction is 2-6 by volume, according to claim 22 the method of. (1) hydrotreating unit, wherein, after the raw material is contacted with hydrogen and a hydrotreating _catalyst, H 2, CH _4, hydrotreated _C ^{2 0} _-C ⁴ 0, hydrotreated naphtha, and hydrotreated A product is obtained;
(2) a catalytic cracking unit, wherein the hydrotreating product is fed to the catalytic cracking according to the method of claim 1, wherein spent catalyst is separated from the reaction product vapor, wherein catalyst, and after being reproduced is stripped, the flow returns to all or a portion of the reaction zone of step (2), and the reaction product vapor is separated, H 2 _/ CH _4, catalytic cracking C ₂ ⁼ -C ₃ ⁼ , catalytic cracking C ₂ ⁰ -C ₃ ⁰ , catalytic cracking C ₄ -C ₅ , catalytic cracking naphtha, LCO, and HCO, where C ₂ ⁼ -C ₃ ⁼ is one of the desired products And the C ₄ -C ₅ is recycled to the catalytic pyrolysis reactor;
(3) Steam cracking unit, wherein the hydrotreated C ₂ ⁰ -C ₄ ⁰ , hydrotreated naphtha, and catalytic cracking C ₂ ⁰ -C ₃ ⁰ are steamed at a temperature of 700 ° C to 1000 ° C. contact, and the reaction product vapor is _separated, H 2 / CH _4, steam cracking _C ^{2 =} _-C ³ =, steam cracking _C ^{2 0} _-C ³ 0, steam cracking _C 4 -C _5, steam cracking naphtha, And fuel oil, where the steam cracking C ₂ ⁼ -C ₃ ⁼ is part of the desired product;
(4) a selective hydrogenation unit, wherein the catalytic cracking naphtha and steam cracking naphtha are first fed to selective hydrogenation to obtain a selectively hydrogenated naphtha;
(5) A solvent extraction unit, wherein the selectively hydrogenated naphtha is fed to solvent extraction to obtain an aromatic hydrocarbon and an extracted raffinate, the aromatic hydrocarbon being the desired product And the extracted raffinate returns to step (3) as one of the ingredients for steam cracking,
For producing light olefins and aromatic hydrocarbons.

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